Reactor and process for solid phase continuous polymerisation of polyethylene terephthalate (pet)

ABSTRACT

A reactor and a process for solid phase continues polymerisation of polyethylene terephthalate (PET) in order to achieve an increase of the intrinsic viscosity (I.V.) of a low molecular weight PET pre-polymers flow. A plurality of fluidised stages in series ( 109; 309 ) are provided and said PET pre-polymers flow is fed into said fluidised stages in series ( 109; 309 ). An inert gas flow is fed either in cross-flow or in counter-current flow with respect to the PET pre-polymers flow.

The present invention relates to a reactor and a process for solid phasecontinuous polymerisation of polyethylene terephthalate (PET).

More exactly, the present invention relates to a reactor and a processfor solid phase continuous polymerisation of polyethylene terephthalate(PET) in order to achieve an increase of the intrinsic viscosity (I.V.)of a low molecular weight PET pre-polymers flow.

Generally, PET manufacturing process includes four steps:esterification, pre-condensation, finishing step and solid phasepolymerisation.

In the past, PET manufacturing process was carried out bytransesterification of dimethyl terephthalate (DMT) with ethylene glycol(EG) while direct esterification of purified terephthalic acid (PTA)with ethylene glycol (EG) is nowadays well-established, whose advantagesmainly reside both in its cost effectiveness and in its higher reactionrate.

Ethylene glycol (EG) and oligomers are present among the by-products atthe end of the reaction.

Ethylene glycol (EG) is usually recycled in feed at the beginning of theprocess, thus increasing the overall yield; the oligomers have to beremoved and then be disposed of or, alternatively, recycled to continueincreasing their polymerisation grade.

It is known that the oligomers in the field of interest, i.e. within thePET manufacturing processes and, more generally, within the polyesterones, are normally molecules having 2, 3 or maximum 4 polymerisationgrade, at gaseous state at high temperature conditions and at liquidstate at temperatures lower than 120÷150° C.

Said oligomers are present at gaseous state at the end of the reaction,but tend to condense on the colder metal surfaces of piping/fittingsthey pass through; once condensed, they tend to produce sticky mixturetogether with finest dusts that are not removed by cyclone separationsystems usually provided downstream the PET manufacturing plants.

Such sticky mixture settles onto the pipes walls and progressivelyreduces the flow areas thereof, until causing the whole system shutdownin case it is not properly removed during maintenance operations.

In consequence of what above described, it becomes evident that theoligomers removal from the gaseous flow exiting the plant is essential.

The conventional melt-phase polymerisation processes routinely use thefirst three steps previously mentioned, i.e. esterification,pre-condensation and finishing step, to produce so-called pre-polymersfor the subsequent solid phase polymerisation.

Solid phase polymerisation is a thermal treatment process that allows toincrease the molecular weight of a polymer at different levels,depending on the kind of final product that it is wanted to be obtained.

It is known that the molecular weight of a polymer can be measured bythe measure of its intrinsic viscosity I.V. (Intrinsic Viscosity).

It is also known that the increase of the molecular weight of a polymercan be achieved by subjecting low molecular weight polymers(pre-polymers), preferably under granular form, to a solid phasecontinuous polymerisation step on fluidised bed. Such process furtherallows to provide polymers with low acetaldehyde content, content thatmust be lower than 1 p.p.m. for the manufacturing of PET bottlesintended for food use.

The pre-polymers fed to the solid phase polymerisation step can be underthe form of completely amorphous pellets or partially crystallisedpastilles.

Amorphous PET pellets are thermally stable only up to the glasstransition temperature, which is about 80° C.; however, the temperatureof the solid phase polymerisation process is higher than 200° C.

By subjecting amorphous PET pellets to a heating phase as well as to atemperature maintaining phase above the glass transition temperature, aprogressive crystallisation of the polymeric bulk is achieved, byhardening the polymer matrix inside pellets themselves, so achievingthermal stability up to temperature values close to the onset (about235° C.) of melting point.

Due to this reason, pre-crystallisers and crystallisers are generallynecessary plant units upstream the conventional solid phasepolymerisation reactor.

The purpose of the crystallisation step prior to solid phasepolymerisation is to prevent the agglomeration of the granules duringthe polymerisation process, especially at the highest temperatures.

As a matter of fact, it is known that in nowadays employed industrialsolid phase polymerisation plants, the granules agglomeration phenomenonfrequently happens; this problem is particularly evident during PETsolid phase continuous moving bed polymerisation for producing bottlesintended for food use, when polymerisation is carried out attemperatures above the amorphous PET pre-polymer glass transitiontemperature, but below the melting point.

According to the conventional processes today available and employed forsolid phase polymerisation, the PET pre-polymer (crystallised orpartially crystallised) is fed at the top of a vertical moving or staticbed reactor for solid phase polymerisation, in which the pre-polymermoves downwards by gravity in contact with an inert gas stream.

According to known prior art, the inert gas primarily functions toremove unwanted by-products, in particular ethylene glycol, acetaldehydeand water vapour, which are generated during polymerisation, while PETgradually moves towards the bottom of the vertical reactor.

In general, there are three important requisites that are to be met forcorrect operation of a continuous solid phase polymerisation process.

First, a steady and uninterrupted flow of PET granules must bemaintained.

As a consequence, it is highly important that agglomeration of PETgranules is avoided, since it would impede the smooth flow of granulesand it would make difficult the discharge of the product from thereactor, thereby causing the plant control losing.

Secondly, a suitable combination of residence time and temperature ofgranules in the reactor is required to achieve the desired molecularweight, which is measurable, as indicated above, in terms of intrinsicviscosity I.V.

Since reaction rate increases with temperature increase, and I.V.increases with residence time increase in the reactor, desired I.V. canbe attained either by using a relatively long residence time, with arelatively low temperature, or a relatively short residence time with arelatively high temperature.

However, the ideal combination of residence time and temperature must bechosen taking into account the first of the requisites indicated above,i.e. the need to maintain a linear flow, thereby avoiding the granulesagglomeration phenomenon.

Third, the flow regime of PET granules processed inside the solid phasepolymerisation reactor, must be as close as possible to the idealplug-flow behaviour; this way all PET granules passing through thereactor undergo the same conditions of treatment and a narrow molecularweight distribution in the obtained product and, more in general, anarrow distribution of polymerised granules final attributes, which is akey factor for the required performance during subsequent high I.V. PETproduct processing/converting steps, will be achieved.

As regards the first requisite, that is the need to avoid the PETgranules agglomeration, it is to be said that this phenomenon is mainlyaffected by temperature, granules size, bed height, granules flowvelocity within the reactor and PET morphology.

PET granules, initially moving freely in a moving bed can stick and clotif, for instance, temperature or bed height increases or ifcross-sectional velocity decreases.

At solid phase polymerisation conditions, PET is only partiallycrystallised and, as a consequence, such a PET is not a rigid body, butit is rather a slightly sticky body.

Since the PET tendency to become sticky increases with temperatureincrease, the agglomeration tendency also increases with temperatureincrease.

A fixed bed of PET granules held motionless inside a solid phasepolymerisation vertical, cylindrical reactor is taken into account.

Under these conditions, at polymerisation temperature and under pressuredue to the weight of the PET granules bed, granules to be polymerisedcreep into one another at contact points and, in time, polymer granuleswill tend to agglomerate and form larger lumps.

The most effective way to prevent lumping is to constantly renew theintergranular contact areas, so that polymer granules do not have achance to creep into one another.

This purpose is achieved by maintaining a constant flow of polymergranules at sufficiently high velocity.

Since agglomeration tendency increases with the increase of the specificsurface area (area per unit mass) or, more precisely, with the increaseof the specific contact area of polymer granules, it also increases withthe decrease of the size of polymer granules.

A reduced granules size contributes to accelerate the polymerisationprocess, however, on the other hand, the agglomeration tendency ofpolymer granules increases.

In presence of small size granules it is therefore required tocounterbalance the increased agglomeration tendency with a temperaturedecrease, which, on the other hand, brings the final values of theprocess rate back to the typical ones for larger size granules treatedat a higher temperature.

Furthermore, if the particle size is reduced below certain limits,agglomeration occurs practically at any temperature.

In a static or moving bed situation, the compaction pressure undergoneby polymer granules is approximately proportional to the weight of thepolymer granules inside the bed which, in its turn, is proportional tothe bed height above the granules; therefore, polymer agglomerationtendency is highest at the bottom of the bed and lowest at the top.

From what above mentioned it results that polymer lumps start generallyto form near the bottom of the bed; for this reason there is a practicallimit on the bed height of a solid phase polymerisation reactor.

At a sufficiently high flow velocity, polymer granules change theirpositions relative to each other (by sliding, for instance) and lumpsformation is thereby prevented.

Since the change rate of polymer granules contact areas and thereduction of bed bulk density increase with the increase of the granulesvelocity, polymer agglomeration tendency inside the reactor decreaseswith the increase of the granules cross-sectional velocity.

For each combination of reactor temperature, bed height, and particlesize, there exists a minimum granules velocity necessary to preventagglomeration.

For each given size and shape of polymer granules, the minimum velocityto prevent agglomeration increases with the increase of temperature andbed height.

In case polymerisation temperature increases or in case bed heightincreases, a higher velocity has to be used; in case, for instance, ofcommercial scale vertical reactors, with outputs up to 300 metric tonsper day and which are conventionally 18 to 22 meters high, a velocity ofat least 2 meters per hour is generally required.

A well designed solid phase polymerisation commercial scale plant mustbe capable of continuously producing outputs having intrinsic viscosityI.V. in compliance with the required specification at a sufficientlyhigh throughput.

The currently used plants employ single or multiple vertical cylindricalreactors 10 to 30 meters high; in those plants the reactor is operatedat a temperature of 200 to 230° C. and at a granules moving velocity of1.00 to 2.52 meters per hour. Within the above-mentioned ranges oftemperature, bed height, and granules velocity, the choice to achieve aproduct with the desired I.V. can be made.

These conventional currently used plants allow to produce PET having anI.V. of 0.72 to 0.86 dl/g, using PET pre-polymers with an I.V. of 0.55to 0.65 dl/g; these conventional plants can increase polymer I.V. byabout 0.12-0.25 dl/g.

Amorphous PET granules have intrinsic viscosity values generallycomprised in the range 0.57÷0.62 dl/g; the reaction time necessary toachieve final I.V. value in the range 0.72÷0.85 dl/g, required for themost bottle manufacturing applications, is of 12÷18 hours.

Usually, PET I.V. is brought to the above-mentioned values, commerciallyrequired for bottle manufacturing, through polymerisation processescarried out in continue solid phase vertical reactors, in which solidbed of PET granules moves downwards by gravity.

For some specific applications, for instance PET pre-polymerspolymerisation for standard bottle manufacturing, characterised byinitial I.V. values of 0.25-0.45 dl/g, it is necessary to increase saidI.V. by more than 0.25 dl/g.

This result is hardly achievable and often it is not achievable in aconventional plant using said vertical reactors.

In a conventional process, there are two ways to raise the product I.V.;namely, to increase the reactor temperature or to increase the granulesresidence time inside the reactor. The residence time inside the reactoris constrained by bed height and by granules velocity; it can beincreased either by increasing bed height or by decreasing granulesvelocity.

Increasing the reactor diameter allows an increase in the throughputrate, but not in residence time at constant granules velocity.

On the other hand, if reactor temperature is raised to increase thefinal product I.V., polymer agglomeration tendency will consequentlyincrease.

To prevent polymer agglomeration, bed height has to be decreased orgranules velocity has to be increased. However, both these modificationsreduce residence time inside the reactor and nullify the effect of thetemperature increase. Alternatively, if the residence time inside thereactor is increased either by increasing the bed height (assuming thereactor is sufficiently tall) or by reducing the granules velocity, anincrease of the polymer agglomeration tendency is caused.

To prevent the agglomeration phenomenon, the reactor temperature must belowered and this once again nullifies the effect of the residence timeincrease on the product final I.V.

These constraints limit the ability of conventional plants, which usesingle or multiple vertical reactors, to increase the polymer intrinsicviscosity I.V.

A similar situation is encountered when a commercial scale plant with acapacity above 360 metric tons per day has to be designed forconventional continuous solid phase polymerisation processes.

In fact, in a conventional process, there are two ways to reach highplant production capacity: again by increasing the reactor temperatureor by increasing the product volume (“hold-up”) in the reactor.

As far as drawbacks due to the temperature increase are concerned, thesame above described issues have to be considered.

On the other hand, the product volume (“hold-up”) of PET granules in thereactor is constrained by bed height, reactor diameter and granulesvelocity.

If the product volume (“hold-up”) is increased either by increasing bedheight or reactor diameter, or by decreasing granules velocity, polymeragglomeration tendency will increase too.

Thus, these constraints limit the maximum capacity of conventional solidphase polymerisation processes, which. use one or more verticalcylindrical reactors.

Nowadays, growing PET demand has given rise to a need for solid phasepolymerisation processes by means of which it is possible to achieve ahigher increase of PET molecular weight and a higher productioncapacity, typically higher than 300 metric tons per day on single plant.

One of the drawbacks of the plants according to the known prior art isdue to the considerable vertical size of the plant structure due to thepresence of the pre-crystallisation unit, which is stacked onto thecrystallisation unit, stacked in its turn onto the solid phasepolymerisation reactor.

A second drawback is represented by the long residence time, inside theconventional reactor, required to achieve the desired I.V. increasestarting from amorphous pre-polymers; for example, to get from an I.V.of 0.60 dl/g to that of the final product, which is equal to 0.80 dl/gin the case of application in the manufacturing of standard bottlesintended for food use, an average residence time equal to 15 hours isneeded.

Such drawback is substantially due to two reasons: the fact that thesolid phase polymerisation kinetic is linked to the reaction gaseousproducts diffusivity from the granules core to their outside and thefact that the pre-crystallisation and crystallisation steps, which aremandatory to secure enough flowability to the granules in a conventionalmoving bed solid phase polymerisation reactor, limit the polymericmatrix portion able to react, being the only non-crystalline portion tobe involved in the reaction. A known solution to the above-mentionedproblems consists in the use of PET sand of spherical granules having asufficiently small size (typically 100÷200 μm) and of a great amount ofinert gas in the reactor.

Thanks to these expedients the pre-crystallisation and crystallisationsteps have been avoided, the reaction time has been reduced and thesolid phase polymerisation reactor acquires lower size.

The aforesaid solution is nevertheless not optimal from the point ofview of cost, of reactor size, of the achievable I.V., of I.V.distribution and of the reaction time.

A first purpose of the present invention is therefore to provide animproved reactor for solid phase continuous polymerisation ofpolyethylene terephthalate (PET) in order to achieve an increase of theintrinsic viscosity I.V. that allows to overcome the constraints of theprocesses known so far, that is to reduce the process time and thereactor size.

A further purpose of the invention is therefore to provide a reactor anda process for PET solid phase polymerisation that allows to achieve highproduction capacities.

In the solid phase polymerisation plants also the gas flow-rate has tobe sufficient to effectively remove the reaction by-products; as amatter of fact, a gas excess results in higher costs both for its supplyand for its regeneration and disposal.

Therefore, a further purpose of the invention is to provide a reactorand a process for solid phase polymerisation that allows to reduce thecosts due to the gas employment.

These and other purposes are achieved with the reactor and the processas claimed in the attached claims.

Advantageously, thanks to the provision of a plurality of fluidisedstages to carry out the solid phase continuous polymerisation of a lowmolecular weight PET pre-polymers flow, a polymerisation up to thedesired I.V. value in a moderate reaction time by using a compactreactor is achieved.

Advantageously, the reactor and the process according to the inventionfurther allow to avoid unwanted agglomeration phenomena and other sideeffects even achieving higher molecular weight increases of the treatedPET when compared with the ones achievable with the conventionalprocesses of the known prior art.

The reactor and the process according to the invention further allows toachieve a high degree of plug-flow (“plug-flow”) and, consequently ahigh homogeneity and uniformity of the final product.

Advantageously, moreover, the reactor according to the invention allowsto achieve higher production capacities when compared with the plantsexploiting the conventional processes.

Always according to the invention is furthermore possible toadvantageously achieve a reduction of energy consumption, thanks to thedecreased overall process AP required for the gas with respect toconventional processes.

The optimal number of fluidised stages was ascertained experimentally,by employing both a model formed by single fluidised bed reactorsarranged in series, and a fluidised reactor model within which aplurality of fluidised stages is generated.

The invention will be now described more in detail with particularreference to the attached drawings, provided by way of not limitingexample, wherein:

FIG. 1A is a diagram of a first embodiment of the invention;

FIG. 1B is a diagram according to an alternative form of the embodimentof the invention shown in FIG. 1A;

FIGS. 2A to 2E show the curves of some quantities that are significantfor achieving the purposes of the invention;

FIG. 3A is a diagram of a second embodiment of the invention;

FIG. 3B is a diagram according to an alternative form of the embodimentof the invention shown in FIG. 3A.

With reference to FIG. 1A a reactor 101 comprising a horizontallyarranged casing 103, having a substantially parallelepiped shape, isshown.

A first feeding line 111, equipped with a device 113 (for example, arotating volumetric distributing apparatus) to regulate the flow-rate offed amorphous PET pre-polymer and to prevent gas leakage, is provided tofeed a continuous low molecular weight PET pre-polymers flow (preferablya PET sand having granules with size in the range 100÷200 μm) to the topof reactor 101.

A second gas feeding line 121 is provided to feed through supply valves123 an inert gas flow, preferably nitrogen, from the bottom into thereactor 101, in cross-flow with respect to the pre-polymer flow insidethe reactor 101.

A circuit 131 connected to the reactor 101 is provided to purify the gasand to recover pre-polymer particles by means of a proper separator 133,such as a cyclone; solid particles dragged by the gas exiting thereactor 101 and recovered by means of the separator 133, are thenrecycled to the reactor 101.

According to an alternative form, shown in FIG. 1B, the gaseous flowcontaining, among others, ethylene glycol and oligomers, is forced topreventively pass through an appropriate separation device 135, such asfor example a condensation separator, wherein ethylene glycol andoligomers are recovered at the liquid state and then recycled to theesterification reactor RE that is located upstream the manufacturingplant to carry out the primary reaction of the PET manufacturingprocess.

The gas without solid particles leaving the separator 133 from its head,is conveyed to a following purification unit (not shown) wherein all theresidual hydrocarbons that are present in the gaseous flow are removedby means of known techniques, for example the catalytic oxidation; saidgas flow is then subjected to drying, for example for adsorption onmolecular sieves, and then recycled to the solid phase polymerisationreactor 101 as fluidising and stripping means of the reaction gaseousby-products.

A third discharging line 141, equipped with a device 143 (for example, arotating volumetric distributing apparatus) to discharge the PET aftersolid phase polymerisation and to prevent gas leakage, is provided todischarge the polymerised product by being inferiorly connected to thebottom of the reactor 101.

Advantageously, according to the invention, inferior vertical walls 105a, secured to the base of the reactor 101, and superior vertical walls105 b, secured to the ceiling of the reactor 101, are provided insidethe reactor 101.

Said walls 105 a and 105 b are preferably formed from metal platessecured, for example by welding, respectively to the base 107 a of thecasing 103 and to the ceiling 107 b of the casing 103.

Moreover, said walls 105 a have decreasing height in the polymer flowadvancing direction, while said walls 105 b have increasing height inthe same direction in order to prevent that the polymer flow could“skip” one or more fluidised stages 109.

Said walls 105 a and 105 b are further spaced so as to create betweenthem corresponding fluidised stages 109, which are generated thanks tothe action of the opposite direction gas flow coming from the holedbottom 104 of the casing 103.

Advantageously, according to the invention, said supply valves 123 arein such a number and arranged in such a way to generate a sufficient gasflow in correspondence with each fluidised stage 109.

Moreover, each supply valve 123 or each supply valves 123 set associatedto the same fluidised stage, is equipped with a heating device 125suitable for bringing at the desired temperature the inert gas flowingthrough the bottom 104 of the reactor 101.

This way, still according to the invention, each fluidised stage 109 canreceive inert gas at the optimal temperature chosen in dependence of thepolymerisation degree reached inside the reactor 101 so achieving anon-isothermal process.

According to the invention, the pre-polymers flow, coming from thefeeding line 111, passes through a plurality of fluidised stages 109,prior to reach the discharge line 141 located downstream the lastfluidised stage.

During crossing reactor 101, pre-polymers are conveniently polymerisedby achieving the desired intrinsic viscosity I.V.

According to the invention, the intrinsic viscosity value attained atthe end of the reaction depends on the number of the fluidised stages.

As it will become evident from the following description, said fluidisedstages are preferably in the number of five.

In order to optimise the number of fluidised stages to be created insidethe reactor, some simulations have been carried out by employing asmodel a plant formed by a plurality of single fluidised bed reactorsarranged in series.

With reference to FIGS. 2A and 2B, the trend of the average I.V. finalvalue with respect to the number of reactors of said plant model isshown.

FIG. 2A refers to the case wherein pre-polymers having an initial I.V.of 0.30 dl/g are used, with overall residence time of three hours andwith reaction temperature inside each reactor of 210° C. (curve T1),22020 C. (curve T2) and 230° C. (curve T3).

As it will be noticed by examining the trend of the three curves T1, T2and T3, the number of single fluidised bed reactors beyond which thefinal I.V. increase is no more significant is about five.

FIG. 2B refers to the case wherein the overall residence time is of fivehours and the steady reaction temperature inside each reactor is of 220°C., by using pre-polymers having an initial I.V. of 0.20 dl/g (curveV1), 0.26 dl/g (curve V2) and 0.30 dl/g (curve V3).

As it will be noticed by examining the trend of the three curves V1, V2and V3, the number of single fluidised bed reactors beyond which thefinal I.V. increase is no more significant is about five.

With reference to FIG. 2C, the trend of the average I.V. final valuewith respect to the reaction time required to attain the desiredpolymerisation is shown.

FIG. 2C refers to the case wherein pre-polymers having an initial I.V.of 0.30 dl/g are used, with reaction temperature of 230° C. and with anumber of fluidised stages of one (curve S1), five (curve S2) and twenty(curve S3).

As it will be noticed by examining the trend of the three curves S1, S2and S3, the number of fluidised stages beyond which the reaction timereduction is no more significant is about five.

FIG. 2D refers to the case wherein pre-polymers having an initial I.V.of 0.30 dl/g are used, with a plant having five fluidised stages andwith reaction temperature of 210° C. (curve T4), 220° C. (curve T5) and230° C. (curve T6).

As it will be noticed by examining the trend of the three curves T4, T5and T6, in a plant provided with five fluidised stages, meaning toachieve a final intrinsic viscosity I.V. value of about 0.84 dl/g, themore the reaction temperature is high the more the reaction timereduction is significant.

FIG. 2E refers to the case wherein a plant having five fluidised stagesis used, with temperature of 230° C. and by using pre-polymers having aninitial I.V. of 0.20 dl/g (curve V4) and 0.30 dl/g (curve V5).

As it will be noticed by examining the trend of the two curves V4 andV5, in a plant provided with five fluidised stages, meaning to achieve afinal intrinsic viscosity I.V. value of about 0.84 dl/g, a moderateincrease of the pre-polymer initial I.V. results in a considerablereduction of the reaction time.

With reference to FIG. 3A a second embodiment of the invention is shown,wherein a reactor 301 comprising a vertically arranged casing 303,having a substantially cylindrical shape, is provided.

A first feeding line 311, equipped with a device 313 (for example, arotating volumetric distributing apparatus) to regulate the flow-rate offed amorphous PET pre-polymer and to prevent gas leakage, is provided tofeed, from the top of reactor 301, a continuous low molecular weight PETpre-polymers flow, preferably a PET sand having granules with size inthe range 100÷200 μm.

A second gas feeding line 321 is provided to feed through supply valves323 an inert gas flow, preferably nitrogen, from the bottom into thereactor 301, in counter-current flow with respect to the descendingpre-polymer flow inside the reactor 301.

A circuit 331 connected to the top of the reactor 301 is provided topurify the gas and to recover pre-polymer particles by means of a properseparator 333, such as a cyclone; solid particles dragged by the gasexiting the reactor 301 and recovered by means of the separator 333, arethen recycled to the reactor 301.

According to an alternative form, shown in FIG. 3B, the gaseous flowcontaining, among others, ethylene glycol and oligomers, is forced topreventively pass through an appropriate separation device 335, such asfor example a condensation separator, wherein ethylene glycol andoligomers are recovered at the liquid state and then recycled to theesterification reactor RE that is located upstream the manufacturingplant to carry out the primary reaction of the PET manufacturingprocess.

The gas without solid particles leaving the separator 333 from its head,is conveyed to a following purification unit (not shown) wherein all theresidual hydrocarbons that are present in the gaseous flow are removedby means of known techniques, for example the catalytic oxidation; saidgas flow is then subjected to drying, for example for adsorption onmolecular sieves, and then recycled to the solid phase polymerisationreactor 301 as fluidising and stripping means of the reaction gaseousby-products.

A third discharging line 341, equipped with a device 343 (for example, arotating volumetric distributing apparatus) to discharge the PET aftersolid phase polymerisation and to prevent gas leakage, is provided todischarge the polymerised product by being inferiorly connected to thebottom of the reactor 301.

Advantageously, according to the invention, almost horizontal shelves305 are provided inside the reactor 301.

Said shelves 305 are preferably formed from holed metal plates secured,for example by welding, to the internal wall 307 of the casing 303 so toextend substantially up to the centre line of the chamber created insidethe casing 303.

Said shelves 305 are further preferably alternately arranged so todefine, inside the reactor 301, an obligatory path that forces thepre-polymers flow to pass through at least two or more plates on which,thanks to the action of the opposite direction gas flow, correspondingfluidised stages 309 are thus generated.

According to the invention, the pre-polymers flow, coming from thefeeding line 311, passes through a plurality of fluidised stages 309,prior to reach the reactor bottom and to be removed through thedischarge line 341.

During crossing reactor 301, pre-polymers are conveniently polymerisedby achieving the desired intrinsic viscosity I.V.

According to the invention, the intrinsic viscosity value attained atthe end of the reaction depends on the number of the fluidised stages.

As it is evident from the preceding description, said fluidised stagesare preferably in the number of five.

With reference to the embodiment of FIG. 1A, three carrying out examplesare hereinafter described.

EXAMPLE 1

Some tests with “PET sand” have been carried out in a pilot plantcomprising a fluidised reactor 101 having five stages in series, ofparallelepiped shape.

Downstream the reactor, after the device 143 that in this example was agas-tight rotating volumetric distributing apparatus device, a coolingunit of the (single stage) fluidised bed type was provided wherein airhas been used as heat transferring and fluidising medium.

The solid phase polymerisation reactor 101, as already said, is amultistage (five stages arranged in series) fluidised reactor havingparallelepiped shape, the base being 1,400 mm×500 mm (gas distributionplate total surface) and the height being 3,200 mm, wherein the solid isfed at a end of the long side and the same progressively reacts byoverflowing from a stage to the following one according to modescomparable to that of a fluid, and wherein the fluidising, heattransferring and gaseous reaction by-products removing gas is fed incross-flow with respect to the solid phase.

The five stages with same volumes, completely independent from the pointof view of both the gas fluid-dynamics and the solid one, are obtainedby means of dividing walls 105 a having decreasing height (respectively2,000, 1,900, 1,800, 1,700 and 1,600 mm).

To prevent that the “PET sand” could “skip” one or more stages, close tothe centre line of each stage a descending. deflector (not shown) hasbeen inserted, having an aperture in the top portion, said aperturebeing finalised to allow the removal of the gas exiting the singlestages avoiding an acceleration of the gas itself close to theoverflows.

Each stage receives from the bottom the hot gas, through a distributionplate formed by a sintered steel plate, whose porosity is of 6.2% andwhich assures a ΔP of about 50 mbar, so to guarantee a uniformdistribution of the gas itself.

Each stage, as aforesaid, has a dedicated gas feeding circuit that isformed by a butterfly valve for flow-rate regulation, an electric heaterfor gas temperature control and a “settling chamber” onto which theabove-mentioned gas distribution sintered plate is placed.

After having performed its numerous functions (heat transfer,fluidisation and reaction gaseous by-products removal) the gas isdischarged from the reactor 101 passing through at first an inertialseparator (for example, of the multi-plate type) and then a centrifugalseparator (for example, a cyclone) 133, both suitable for separating thepossible dragged “PET sand” particles.

The employed gas velocity is of 3 times with respect to the minimumfluidisation velocity for the first stage and 2 times with respect tothe minimum fluidisation velocity for the stages from the 2^(nd) to the5^(th), this to create a higher vacuum degree in the first fluidisedstage so to oppose the agglomeration tendency of. the fed cold amorphous“PET sand”, while for the stages from the 2^(nd) to the 5^(th), havingthe “PET sand” crystallinity values gradually increasing (as aconsequence of the maintaining at temperature), said agglomeration riskis reduced.

The gas is nitrogen (purity: 99.99%) with a dew point of −45° C. For thesolid phase polycondensation test a “PET sand”, formed by particles of200 μm size (more precisely: 99.9%<220 μm and 0.1%<180 μm) of amorphousPET containing 2.0% by weight of isophthalic acid and having a meltingpoint of 251° C., has been employed.

The PET granules flow was of 250 kg/h; the overall hold-up of solid inthe five stages (at the fluidisation conditions above-mentioned) wasmeasured to be 360 kg (weigh of the “PET sand” mass contained in thefive stages when stopping the gas feeding): it results an averageresidence time of the solid equal to 1 hour and 25 minutes (360 kg/250kg/h). The pre-polymer starting intrinsic viscosity was 0.30±0.005 dl/g;the acetaldehyde content was 47 p.p.m.; the PET granules temperature atthe entrance of the reactor was 23° C. The temperature of the gas fed tothe five stages was 228±0.5° C.

The same “PET sand”, after solid phase polymerisation, at the exit ofthe reactor had a final intrinsic viscosity of 0.82±0.01 dl/g.

EXAMPLE 2

In this second carrying out example of the process according to theinvention, all the conditions applied in the first example have beenretained, with the exception of the fed gas temperature, which has beenset as follows:

-   1^(st) stage: 228±0.5° C.;-   2^(nd) stage: 229±0.5° C.;-   3^(rd) stage: 229±0.5° C.;-   4^(th) stage: 230±0.5° C.;-   5^(th) stage: 230±0.5° C.

The “PET sand”, after solid phase polymerisation, at the exit of thereactor had a final intrinsic viscosity of 0.854±0.01 dl/g.

EXAMPLE 3

In this third carrying out example of the process according to theinvention, all the conditions applied in the first example have beenretained, with the exception of the intrinsic viscosity of the fedamorphous “PET sand” and of the temperature of the gas fed to the fivestages.

In particular, amorphous “PET sands” have been used having the followingdifferent intrinsic viscosity parameters: 0.26±0.005 dl/g; 0.36±0.005dl/g; 0.45±0.005 dl/g; 0.58±0.005 dl/g. In all these tests the gasfeeding temperature (gas temperature equal for the five stages) has beenmade to change, so as to achieve a final intrinsic viscosity of 0.82dl/g for the “PET sand” after polymerisation.

Obviously, not having changed the feeding flow-rate of both the solidand the gas, the average residence time of the solid remained of 1 hourand 25 minutes.

The gas temperatures, to achieve the same solid phase polymerised PETfinal intrinsic viscosity of 0.82 dl/g for starting from the fourdifferent “PET sands”, have been respectively: 229±0.5° C.; 225±0.5° C.;218±0.5° C.; 211±0.5° C.

In each of the checked conditions, the final intrinsic viscosityvariation have been noticed to be of ±0.01 dl/g, corresponding to anincrease of the final intrinsic viscosity variation due to solid phasepolymerisation of ±0.005 dl/g, this confirming the good “plug-flow”behaviour achieved on the solid phase in the serial multistage process.

It has to be noticed that in the conventional solid phase polymerisationprocesses, such as those carried out inside a cylindrical, with verticalaxis, moving bed reactor, the increase of the final intrinsic viscosityvariation due to solid phase polymerisation is of 0.01±0.015 dl/g.

The same test has been also performed in a conventional, cylindrical,vertical, moving bed reactor having an internal diameter of 1.6 meters,bed height of 8 meters and “PET sand” downhill velocity of 0.42meters/hour.

The “PET sand”, in this case, was previously crystallised to a degree ofcrystallinity X_(c)=38% so as to give proper flowability to thesubsequent moving bed reactor, by heating and maintaining a temperatureof 200° C. for 5 minutes.

All the tests performed with this plant arrangement have highlighted aprocess threshold value of 216±0.5° C., beyond which irreversibleagglomeration phenomena among particles started to occur, this resultingin a progressive loss of flowability.

Only with “PET sand” having initial I.V. of 0.45±0.005 dl/g and with theone having initial I.V. of 0.58±0.005 dl/g it was possible to solidphase polymerise, achieving a final I.V. of 0.82 dl/g.

In the first case the residence time was of 16 hours at a temperature of215±0.5° C. and in the second one of 9.5 hours at a temperature of215±0.5° C.

The exposed example clearly shows that a conventional, cylindrical,vertical, moving bed reactor has a threshold maximum temperature as wellas a limit to the intrinsic viscosity increase achievable by keepingunchanged the reactor, the bed height and velocity.

In this specific case, the allowed maximum temperature of the reactorwas of about 215° C. and the achievable maximum intrinsic viscosityincrease was of about 0.37 dl/g (0.82 final−0.45 initial) with a bedreactor height of 8 meters and a velocity of 0.42 meters/hour.

From what above disclosed, it clearly results that the process accordingto the invention allows to achieve a higher PET molecular weightincrease as well as to operate at a temperature significantly higherthan those previously employed with the conventional moving bedprocesses, without agglomeration phenomena and other unwanted sideeffects and achieving a good result in terms of uniformity of theproduct final characteristics, thanks to the “plug-flow” behaviour onthe solid phase in the plurality of fluidised stages in series.

Furthermore the invention will be advantageously applicable to anypolyester which can be solid phase polymerised. The most commonpolyesters suitable for use in the invention have at least about 75 molepercent of their acid moieties provided by an aromatic dicarboxylicacid, such as terephthalic acid, isophthalic acid, or a naphthalenicdicarboxylic acid (preferably 2,6-) with the diol moieties provided byglycols such as ethylene glycol, butylene glycol, 1,4-dimethylolcyclohexane and the like or aromatic diols such as hydroquinone andcatechol. Said polyesters can contain other dicarboxylic acids such asadipic acid, isophthalic acid, sebacic acid, and the like. Polyethyleneterephthalate, polyethylene isophthalate, polyethylene naphthalate, andpolybutylene terephthalate homopolymers are representative examples ofsuch polyesters.

Blends of various polyesters can also be solid phase polymerised in theprocess according to the invention. The polyester pre-polymers(amorphous starting polyesters) utilised in this invention can be madein any manner but are typically prepared by conventional melt phasepolymerisation techniques. These polyester pre-polymers have an initialstarting IV of at least about 0.2 dl/g as measured in a 60:40 (byweight): phenol ÷1,1,2,2,-tetrachloroethane solvent system at atemperature of 30° C. The rate at which polyethylene terephthalatepre-polymer can be solid state polymerised also depends on the carboxylend group (i.e.—COOH) content of the pre-polymer. Generally,pre-polymers having a carboxyl end group content within the range ofabout 18% to about 40% achieve maximum solid state polymerisation rates.It is preferred for such pre-polymers to have a carboxyl end groupcontent within the range of about 24% to 33% (see for example U.S. Pat.No. 4,238,593). Suitable polyester pre-polymers which can be solid statepolymerised using my invention are comprised of one or more diacidcomponents and one or more diol components. The diacid component in thepolyesters are normally alkyl dicarboxylic acids which contain from 4 to36 carbon atoms, diesters of alkyl dicarboxylic acids which contain from6 to 38 carbon atoms, aryl dicarboxylic acids which contain from 8 to 20carbon atoms, diesters of aryl dicarboxylic acids which contain from 10to 2.2 carbon atoms, alkyl substituted aryl dicarboxylic acids whichcontain from 9 to 22 carbon atoms, or diesters of alkyl substituted aryldicarboxylic acids which contain from 11 to 22 carbon atoms.

The preferred alkyl dicarboxylic acids will contain from 4 to 12 carbonatoms. Some representative examples of such alkyl dicarboxylic acidsinclude glutaric acid, adipic acid, pimelic acid, and the like. Thepreferred diesters of alkyl dicarboxylic acids will contain from 6 to 12carbon atoms. A representative example of such a diester of an alkyldicarboxylic acid is azelaic acid. The preferred aryl dicarboxylic acidscontain from 8 to 16 carbon atoms. Some representative examples of aryldicarboxylic acids are terephthalic acid, isophthalic acid, andorthophthalic acid.

The preferred diesters of aryl dicarboxylic acids contain from 10 to 18carbon atoms. Some representative examples of diesters of aryldicarboxylic acids include diethyl terephthalate, diethyl isophthalate,diethyl orthophthalate, dimethyl naphthalate, diethyl naphthalate andthe like. The preferred alkyl substituted aryl dicarboxylic acidscontain from 9 to 16 carbon atoms and the preferred diesters of alkylsubstituted aryl dicarboxylic acids contain from 11 to 15 carbon atoms.

The diol component of the polyester pre-polymers is normally comprisedof glycols containing from 2 to 12 carbon atoms, glycol etherscontaining from 4 to 12 carbon atoms, and polyether glycols having thestructural formula HO-(A-O)_(n)—H wherein A is an alkylene groupcontaining from 2 to 6 carbon atoms and wherein n is an integer from 2to 400. Generally, such polyether glycols will have a molecular weightof 400 to about 4000. Preferred glycols normally contain from 2 to 8carbon atoms and more preferably from 4 to 8 carbon atoms. Somerepresentative examples of glycols that can be utilised as the diolcomponent include ethylene glycol, 1,3-propylene glycol, 1,2-propyleneglycol, 2,2-diethyl-1,3-propane diol, 2,2-dimethyl-1,3-propane diol,2-butyl-1,3-propane diol, 2-ethyl-2-isobutyl-1,3-propane diol,1,3-butane diol, 1,4-butane diol, 1,5-pentane diol, 1,6-hexane diol,2,2,4-trimethyl-1,6-hexane diol, 1,3-cyclohexane dimethanol,1,4-cyclohexane dimethanol, and 2,2,4,4-tetramethyl-1,3-cyclobutanediol. Some representative examples of polyether glycols that can be usedinclude polytetramethylene glycol and polyethylene glycol.

Branched polyester pre-polymers can also be solid state polymerised inthe process of the present invention. Such branched polyesters normallycontain branching agents which have three or more functional groups andpreferably three or four functional groups. Reactive functional groupscan be carboxyl groups or aliphatic hydroxyl groups. The branching agentutilised in such branched polyesters can optionally contain bothcarboxyl groups and hydroxyl groups. Examples of acidic branching agentsinclude trimesic acid, trimellitic acid, pyromellitic acid, butanetetracarboxylic acid, naphthalene tricarboxylic acids, andcyclohexane-1,3,5-tricarboxylic acids. Some representative examples ofhydroxyl branching agents (polyols) include glycerin, trimethylolpropane, pentaerythritol, dipentaerythritol, 1,2,6-hexane triol, and1,3,5-trimethylol benzene. Generally, from 0 to 3 percent of a polyolcontaining from 3 to 12 carbon atoms will be used as the branching agent(based upon the total diol component).

High strength polyesters which utilise at least one bis-hydroxyalkylpyromellitic diimide in their diol component can also be solid statepolymerised. The diol component in these polyesters will normallycontain from 5 to 50 mole percent of one or more bis-hydroxyalkylpyromellitic diimides and will preferably be comprised of from 10 to 25mole percent of at least one bis-hydroxyalkyl pyromellitic diimide. Theremaining portion of the diol component is comprised of additionalcopolymerisable diols.

1-61. (canceled)
 62. A reactor for solid phase continuous polymerisationof polyethylene terephthalate (PET), comprising: a casing; a feedingline to feed a low molecular weight PET pre-polymers flow into saidreactor; a gas line to feed, through supply valves, a gas into saidreactor; a discharging line inferiorly connected to the bottom of thereactor to discharge the polymerised product; a circuit connected to thereactor to purify the gas and to recover pre-polymer particles by meansof a proper separator; wherein, inside said reactor, means are providedto generate a plurality of fluidised stages in series to cause anincrease of the intrinsic viscosity (I.V.) of said PET pre-polymersflow.
 63. The reactor according to claim 62, wherein said casing of saidreactor has a substantially parallelpiped shape and it is horizontallyarranged.
 64. The reactor according to claim 63, wherein said meanscomprise a plurality of inferior vertical walls secured to the base ofsaid reactor and a plurality of superior vertical walls secured to theceiling of said reactor said fluidised stages being generated betweensaid inferior walls.
 65. The reactor according to claim 64, wherein saidsupply valves are in such a number and arranged in such a way as togenerate a sufficient gas flow in correspondence with each fluidisedstage.
 66. The reactor according to claim 65, wherein said supply valvesassociated to the same fluidised stage are equipped with heating devicessuitable for bringing to the desired temperature an inert gas flowingthrough the bottom of the reactor in correspondence with each fluidisedstage, this way achieving a differentiated heating of said fluidisedstages.
 67. The reactor according to claim 67, wherein said casing ofsaid reactor has a substantially cylindrical shape and the casing isvertically arranged.
 68. The reactor according to claim 67, wherein saidmeans to generate a plurality of fluidised stages comprise a pluralityof shelves secured to the internal wall of said casing in correspondencewith which said fluidised stages are generated.
 69. The reactoraccording to claim 68, wherein said shelves, provided inside saidcasing, are secured to said internal wall of said casing so to radiallyextend up to about the centre line of the chamber created inside saidcasing.
 70. The reactor according to claim 69, wherein said shelves areformed of apertured metal plates.
 71. The reactor according to claim 70,wherein said shelves, provided inside said casing, are approximatelyhorizontal.
 72. The reactor according to claim 71, wherein that saidshelves are alternately arranged so to define an obligatory path insidesaid reactor for the pre-polymers descending flow.
 73. The reactoraccording to claim 62, wherein said fluidised stages in series numberfive.
 74. The reactor according to claim 62, wherein said feeding lineis equipped with a device suitable for regulating the flow-rate of fedamorphous PET pre-polymer and to prevent gas leakage.
 75. The reactoraccording to claim 74, wherein said device is a rotating volumetricdistributing apparatus.
 76. The reactor according to claim 62, whereinsaid discharging line is equipped with a device suitable for dischargingthe PET after solid phase polymerisation and to prevent gas leakage. 77.The reactor according to claim 76, wherein said device is a rotatingvolumetric distributing apparatus.
 78. The reactor according to claim62, wherein said circuit further comprises a separation device torecover ethylene glycol and oligomers at the liquid state and then torecycle them upstream of the overall PET manufacturing cycle.
 79. Thereactor according to claim 62, wherein said PET pre-polymers flow has alow initial l.V. value, generally an l.V. value in the range of0.20÷0.45 dl/g.
 80. The reactor according to claim 62, wherein said l.V.increase of said PET pre-polymers flow is in the range of 0.35÷0.65dl/g.
 81. The reactor according to claim 62, wherein said l.V. increaseof said PET pre-polymers flow is ≧0.06 dl/g.
 82. The reactor accordingto claim 80, wherein said l.V. increase of said PET pre-polymers flow is≧0.20 dl/g.
 83. The reactor according to claim 62, wherein said PETpre-polymers flow is a PET sand flow, the sand particle size preferablybeing in the range 60÷300 μm.
 84. The reactor according to claim 83,wherein said PET pre-polymers flow is a PET sand flow, the sand particlesize preferably being in the range 100÷250 μm.
 85. The reactor accordingto claim 62, wherein said PET pre-polymers flow fed into the reactor ismaintained in said reactor at a temperature in the range 200÷235° C. 86.The reactor according to claim 85, wherein said PET pre-polymers flowfed into the reactor is maintained in said reactor at a temperature inthe range 205÷230° C.
 87. The reactor according to claim 62, whereinsaid gas is an inert gas.
 88. The reactor according to claim 62, whereinsaid gas flow inside said reactor is directed in cross-flow or incounter-current flow with respect to the flow of the PET granules thatpass through said reactor.
 89. The reactor according to claim 62,wherein the ratio between the mass of the gas flow that passes throughsaid reactor and the mass of the PET granules in the reactor is >0.62.90. The reactor according to claim 89, wherein the ratio between themass of the gas flow that passes through said reactor and the mass ofthe PET granules in the reactor is >0.9.
 91. The reactor according toclaim 62, wherein said gas is an inert gas or air.
 92. The reactoraccording to claim 91, wherein said gas is air with a dew point <−30° C.93. The reactor according to claim 91, wherein said gas is a mixture ofgases selected from the group consisting of nitrogen, noble gases,carbon dioxide, carbon monoxide and oxygen, and wherein the oxygencontent is <10% by weight.
 94. The reactor according to claim 91,wherein said gas is a mixture of gases chosen from the group consistingof nitrogen, noble gases, carbon dioxide, carbon monoxide and oxygen,and wherein the oxygen content is <6% by weight.
 95. The reactoraccording to claim 62, wherein the gas is recycled to the reactor, afterhaving been purified of the organic impurities, until a level of organicimpurities ≦100 p.p.m. by weight (CH₄ equivalent) has been reached. 96.The reactor according to claim 62, wherein the PET granules have anirregular shape with a volume comprised between 0.05 and 10 mm³.
 97. Thereactor according to claim 62, wherein inside said reactor the polyestergranules are subjected to a solid phase polycondensation and/or dryingand/or crystallisation and/or dealdehydisation.
 98. A process for solidphase continuous polymerisation of polyethylene terephthalate (PET),comprising the steps of: feeding a low molecular weight PET pre-polymersflow into a reactor through a feeding line; feeding a gas into saidreactor through a gas line in cross-flow or in counter-current flow withrespect to said PET pre-polymers flow, carrying out said polymerisationin a plurality of fluidised stages in series generated inside saidreactor to cause an increase of the intrinsic viscosity (I.V.) of saidPET pre-polymers flow.
 99. The process according to claim 98, whereinsaid polymerisation is carried out in a number of fluidised stages inseries of five.
 100. The process according to claim 99, wherein saidpolymerisation is carried out at non-isothermal conditions.
 101. Theprocess according to claim 99, wherein said polymerisation is carriedout at isothermal conditions.
 102. The process according to claim 99,wherein said polymerisation is carried out in a time period of about 2hours.
 103. The process according to claim 98, wherein said PETpre-polymers flow has a low initial l.V. value, generally an l.V. valuein the range of 0.20÷0.45 dl/g.
 104. The process according to claim 98,wherein said I.V. increase of said PET pre-polymers flow is in the rangeof 0.35÷0.65 dl/g.
 105. The process according to claim 98, wherein saidI.V. increase of said PET pre-polymers flow is ≧0.06 dl/g.
 106. Theprocess according to claim 105, wherein said l.V. increase of said PETpre-polymers flow is ≧0.20 dl/g.
 107. The process according to claim 98,wherein said PET pre-polymers flow is a PET sand flow, the sand particlesize being in the range 60÷300 μm.
 108. The process according to claim107, wherein said PET pre-polymers flow is a PET sand flow, the sandparticle size being in the range 100÷250 μm.
 109. The process accordingto claim 98, wherein said PET pre-polymers flow fed into the reactor ismaintained in said reactor at a temperature in the range 200÷235° C.110. The process according to claim 109, wherein said PET pre-polymersflow fed into the reactor is maintained in said reactor at a temperaturein the range 205÷230° C.
 111. The process according to claim 98, whereinsaid gas is an inert gas.
 112. The process according to claim 98,wherein said gas flow inside said reactor is directed in cross-flow orin counter-current flow with respect to the flow of said PET granulesthat pass through said reactor.
 113. The process according to claim 98,wherein the ratio between the mass of the gas flow that passes throughthe reactor and the mass of the PET granules in the reactor is >0.62.114. The process according to claim 113, wherein the ratio between themass of the gas flow that passes through the reactor and the mass of thePET granules in the reactor is >0.9.
 115. The process according to claim98, wherein said gas is an inert gas or air.
 116. The process accordingto claim 115, wherein said gas is air with a dew point <−30° C.
 117. Theprocess according to claim 115, wherein said gas is a mixture of gaseschosen from the group consisting of one or more of nitrogen, noblegases, carbon dioxide, carbon monoxide and oxygen, and wherein theoxygen content is <10% by weight.
 118. The process according to claim115, wherein said gas is a mixture of gases chosen from the groupconsisting of nitrogen, noble gases, carbon dioxide, carbon monoxide andoxygen, and wherein the oxygen content is <6% by weight.
 119. Theprocess according to claim 98, wherein the gas is recycled to thereactor, after having been purified of the organic impurities, until alevel of organic impurities <100 p.p.m. by weight (CH₄ equivalent) hasbeen reached.
 120. The process according to claim 98, wherein the PETgranules have an irregular shape with a volume comprised between 0.05and 10 mm³.
 121. The process according to claim 98, wherein inside saidreactor the polyester granules are subjected to one or more processesselected from the group consisting of solid phase polycondensation,drying, crystallisation and dealdehydisation.
 122. The process accordingto claim 98, wherein ethylene glycol and oligomers present at the end ofsaid polymerisation are recovered at the liquid state in a separationdevice provided in a circuit connected to the reactor and then recycledupstream of the overall PET manufacturing cycle.